Ignition of oxidative coupling of methane while reducing ratio of methane to oxygen

ABSTRACT

Methods of performing a startup of an oxidative coupling of methane reaction to produce C 2 + hydrocarbons are described. The methods can include incrementally varying startup parameters of the oxidative methane reactor and using the feed gas as a coolant such that high C 2 + hydrocarbon selectivity is achieved.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a divisional of U.S. patent application Ser. No.16/483,563 filed Aug. 5, 2019, which is a national phase applicationunder 35 U.S.C. § 371 of International Application No. PCT/IB2018/050714filed Feb. 5, 2018, which claims the benefit of priority of U.S.Provisional Patent Application No. 62/457,119 filed Feb. 9, 2017, all ofwhich are hereby incorporated by reference in their entirety.

BACKGROUND OF THE INVENTION A. Field of the Invention

The invention generally concerns systems and methods for the productionof C₂+ hydrocarbons from methane (CH₄) and oxygen (O₂). In particular,the systems and methods allow for the use of parameters of reactant feedto a reactor to establish and maintain steady state operation.

B. Description of Related Art

Methane can be used to produce ethane and/or ethylene through theoxidative coupling of the methane (OCM) reaction. While extensiveresearch and development has been devoted to this reaction, the reactionlargely remains inefficient on a commercial scale. One of the keychallenges is the high reaction temperature (typically greater than 750°C.) required to make the reaction proceed. The need for such a hightemperature is due to the bond strength (bond dissociation energy) ofthe tetrahedral C—H bonds in methane, which is 104 kcal per mole(kcal/mol). This C—H bond strength makes methane less reactive anddifficult to undergo oxidative conversion to form ethylene.

The oxidative coupling of the methane is represented by the followingequations:2CH₄+O₂→C₂H₄+2H₂O ΔH=−67.4 kcal/mol  (I)2CH₄+½O₂→C₂H₆+H₂O ΔH=−84.6 kcal/mol  (II)As shown in Equations (I) and (II), oxidative conversion of methane toethylene or ethane is exothermic. It should be noted that the heats ofreaction for Equations (I) to (IV) are given per 100593174.1 mole ofoxygen consumed. Excess heat produced from these reactions can pushconversion of methane to carbon monoxide and carbon dioxide rather thanthe desired C₂ hydrocarbon product:CH₄+1.5O₂→CO+2H₂O ΔH=−82.8 kcal/mol  (III)CH₄+2O₂→CO₂+2H₂O ΔH=−95.9 kcal/mol  (IV)The excess heat from the reactions in Equations (III) and (IV) furtherexacerbate this situation, thereby substantially reducing theselectivity of ethylene production when compared with carbon monoxideand carbon dioxide production.

Equations (V) through (VIII) provide an illustration of the chemicalpathway in which the OCM reaction can occur in the presence of acatalyst:O₂+2[*]→2[O],  (V)CH₄+[O]→CH₃+[OH]  (VI)2CH₃→C₂H₆→C₂H₄→CO_(x)  (VII)2[OH]→[O]+[*]+H₂O  (VIII)where * is a vacant catalytic surface site. The coupling of methylradicals in equation (VII) occurs in the gas phase while the formationof C₂H₄ and CO_(x) may either be catalytic or in the gas phase. Theremaining reactions (V, VI, and VIII) occur on the catalyst.

There are two important practical problems that have prevented thedevelopment of a commercially feasible OCM process: (1) the very largeheat of reaction (Equations I-IV); and (2) the very high temperaturerequired to initiate the reaction (typically 700-950° C.). There is nocommercially available liquid heat transfer fluid capable of operationat such high temperature. Consequently, the only way to cool a reactorat this range of temperature is with very inefficient gas phase coolants(air, steam, or ethane, for example). In a cooled multi-tubular fixedbed reactor, the methane conversion must be limited (by the oxygenconcentration in the feed) to less than about 15% in order to avoid arunaway reaction. A runaway reaction is one in which the temperaturerise within the catalyst bed is high enough to damage or deactivate thecatalyst.

Attempts have been made to address the problem of temperature control byusing one or more substantially adiabatic reactors in series with one ormore heat exchangers (for cooling) following each reactor. By way ofexample, U.S. Patent Application Publication No. 2014/0107385 toSchammel et al. describes a system that uses a series of catalytic bedswhere the inlet temperature is less than 600° C. All of the catalyticbeds are kept at the same temperature throughout the reaction process byremoving thermal energy generated during the upstream reactions. Whilethis system can help with the energy input requirements needed for theOCM reaction, it relies on additional materials and a complicatedcatalytic bed design, which can be expensive and difficult to implementon a commercial scale. Moreover, the methane conversion within eachreactor must be limited to avoid runaway reaction and excessive catalysttemperature.

SUMMARY OF THE INVENTION

A solution to the aforementioned problems associated with the OCMreaction has been discovered. Embodiments of the invention utilizesparameters of the feed reactant gas to establish and maintain reactionconditions in the OCM reactor so that the conversion of methane toethane and/or ethylene is maximized while the destruction of thecatalyst by excessive temperatures in the reactor is avoided. Inparticular, in embodiments of the invention, the OCM reaction is ignitedin a manner that avoids transient state at temperatures that woulddestroy the catalyst. Once in an ignited state, the reaction is made tooperate in an autothermal state by supplying feed gas for the oxidativecoupling reaction to the reactor at a rate and at a low enoughtemperature to compensate for the heat of reaction generated in thereactor. In this way, the feed gas serves as a coolant as it is heatedto a higher temperature by the heat generated by the oxidative couplingreaction in the reactor.

In embodiments of the invention, the reactor is operated as an adiabaticreactor with a particular combination of operating parameters (e.g.,feed composition, feed temperature, feed flowrate, or combinationsthereof) so that the reactor is in an ignited, autothermal state. In theautothermal state, the reactor feed is controlled at a very highflowrate and low temperature thus providing the cooling for removing thelarge heat of reaction and enabling high methane conversion (forexample >15%) within a single reactor.

In one instance of the present invention, a method of performing anoxidative coupling of methane reaction to produce C₂+ hydrocarbons isdescribed. The method can include pre-heating a gaseous feed stream,which has an initial methane to oxygen (CH₄:O₂) molar ratio, to atemperature of at least 400° C. (e.g., 400° C. to 750° C., or 550° C. to650° C.). The preheated gaseous feed stream can be introduced into anadiabatic reactor that includes a catalyst bed containing an oxidativecoupling of methane catalyst. The oxidative coupling of methane reactioncan be ignited. During the startup procedures, both the temperature andthe CH₄:O₂ molar ratio of the gaseous feed stream introduced into theadiabatic reactor can be incrementally reduced to a temperature of 10°C. to 350° C. and a final CH₄:O₂ molar ratio of 9:1 to 3:1 such that thesum of the gaseous feed stream temperature and the rise in temperaturewithin the reactor over the start-up period (e.g., 1 minute to 5 hours)is close to the desired operating temperature (e.g., 750° C. to 1100° C.or 850° C. to 950° C.) of the reaction. By “close,” it is meant that thetemperature at the exit of the reactor is within 15%, or within 5% ofthe final operating temperature. The initial CH₄:O₂ molar ratio can be8:1 to 40:1 and the final CH₄:O₂ molar ratio can be 9:1 to 3:1, 8:1 to5:1, or 6:1 to 5:1, preferably 5.5:1, after incremental reduction. Themethod can further include adding 1 mole % to 10 mole % of a gas morereactive than methane to the gaseous feed stream prior to ignition, anddiscontinuing the addition after ignition. In some instances, thegaseous feed stream can have a residence time of 0.1 to 20 millisecondsin the catalyst bed after ignition. In certain embodiments, the gaseousfeed stream can have a residence time of 25 to 1000 milliseconds in thecatalyst bed before and during ignition, and the method further includesdecreasing the residence time to 0.1 to 20 milliseconds after ignition.The preferred residence time depends on the activity of the catalyst.

In another aspect of the present invention, there is disclosed a startupprocess for an OCM reaction that can include: (a) pre-heating a catalystbed of an adiabatic reactor with a heat source; (b) introducing agaseous feed stream comprising methane (CH₄) and oxygen (O₂) having atemperature of less than 350° C. and an initial CH₄:O₂ molar ratio tothe adiabatic reactor; (c) igniting the oxidative coupling of methanereaction; and (d) incrementally reducing, during startup steps (a) to(c), the initial CH₄:O₂ molar ratio of the gaseous feed streamintroduced into the reactor to a final CH₄:O₂ molar ratio of 9:1 to 3:1.The catalyst bed can include an oxidative coupling of methane catalyst.The temperature of the gaseous feed stream introduced into the reactorcan be 150° C. or less and the catalyst bed can be preheated to 400° C.to 700° C., preferably 550° C. to 650° C., more preferably 500° C. to600° C. In certain embodiments, the catalyst bed can be preheated to theignition temperature of the oxidative coupling of methane reaction atthe initial CH₄:O₂ molar ratio and the catalyst bed will ignite shortlyafter the feed mixture is introduced. (The ignition conditions may befound experimentally.)

After startup procedures have been performed and a steady reaction statehas been reached (e.g., usually within 1 minute to 5 hours afterstartup, depending on the thermal capacitance of the reactor andinsulation), a product stream can be produced that includes C₂+hydrocarbons with a selectivity to C₂+ hydrocarbons of 30% to 95%. Thiscan be achieved with using a reactant feed having a temperature of lessthan 150° C., preferably 75° C. to 125° C., or about 100° C. (i.e.,ambient temperature). In some instances, the product stream alsoincludes hydrogen (H₂) and carbon monoxide (CO), carbon dioxide (CO₂),or mixtures thereof.

The OCM catalyst used in the methods of the present invention can be anycatalyst known in the art. (However, the minimum feed temperaturepossible depends on the catalyst activity.) In certain embodiments, thecatalyst can include a metal oxide, a supported metal oxide, a mixedmetal oxide, a supported mixed metal oxide, or any mixture thereof. In aparticular instance, the catalyst can be La₂O₃/CeO₂, SrO/La₂O₃,Yb₂O₃—SrO—CeO₂, Li/MgO, Na₂WO₄—Mn—O/SiO₂, or any combination thereof.

In one aspect of the present invention, 20 embodiments are described.Embodiment 1 is a method of performing an oxidative coupling of methanereaction to produce C₂+ hydrocarbons, the method comprising: (a)preheating a gaseous feed stream to a temperature of at least 400° C.,wherein the gaseous feed stream comprises methane (CH₄) and oxygen (O₂)having an initial CH₄:O₂ molar ratio; (b) introducing the preheatedgaseous feed stream to an adiabatic reactor, wherein the adiabaticreactor includes a catalyst bed comprising an oxidative coupling ofmethane catalyst; (c) igniting the oxidative coupling of methanereaction; and (d) after igniting the oxidative coupling of methanereaction, incrementally reducing both the temperature and the CH₄:O₂molar ratio of the gaseous feed stream introduced into the adiabaticreactor to an operating temperature of 10° C. to 350° C. and a finalCH₄:O₂ molar ratio of 9:1 to 3:1 over a startup period such that, at theoperating temperature, the oxidative coupling of methane reactionremains ignited and the reactor is in an autothermal state. Embodiment 2is the method of embodiment 1, wherein the initial CH₄:O₂ molar ratio is8:1 to 40:1 and the final CH₄:O₂ molar ratio is 9:1 to 3:1, 8:1 to 5:1,or 6:1 to 5:1, preferably 5.5:1 after incremental reduction. Embodiment3 is the method of embodiment 2, wherein the temperature is reduced in 1to 10° C. increments and the molar ratio is reduced in 0.01 to 1 molarincrements. Embodiment 4 is the method of any one of embodiments 1 to 3,wherein a final catalyst operating temperature is 750° C. to 1100° C. or850° C. to 950° C. Embodiment 5 is the method of any one of embodiments1 to 4, wherein the gaseous feed stream is preheated to 400° C. to 750°C., or 500° C. to 600° C. Embodiment 6 is the method of any one ofembodiments 1 to 5, further comprising: adding 1 mol % to 10 mol % of agas more reactive than methane to the gaseous feed stream prior toignition; and discontinuing the addition after ignition. Embodiment 7 isthe method of any one of embodiments 1 to 6, wherein the gaseous feedstream has a residence time of 0.1 to 1000 milliseconds in the catalystbed before and during ignition, and the method further comprisesdecreasing the residence time to 0.1 to 20 milliseconds after ignition.Embodiment 8 is the method of any one of embodiments 1 to 7, furthercomprising continuing the oxidative coupling of methane reaction afterstep (d) to produce a product stream comprising C₂+ hydrocarbons.Embodiment 9 is the method of embodiment 8, wherein the product streamfurther comprises hydrogen (H₂) and carbon monoxide (CO), carbon dioxide(CO₂), water (H₂O) or mixtures thereof. Embodiment 10 is the method ofany one of embodiments 1 to 9, wherein the catalyst comprises a metaloxide, a supported metal oxide, a mixed metal oxide, a supported mixedmetal oxide, or any mixture thereof. Embodiment 11 is the method ofembodiment 10, wherein the catalyst is La₂O₃/CeO₂, SrO/La₂O₃,Yb₂O₃—SrO—CeO₂, Li/MgO, Na₂WO₄—Mn—O/SiO₂, or any combination thereof.Embodiment 12 is the method of any one of embodiments 1 to 11, whereinthe selectivity of C₂+ hydrocarbons is 30% to 95% after ignition.

Embodiment 13 is a startup method for an oxidative coupling of methanereaction to produce C₂+ hydrocarbons, the method comprising: (a)preheating a catalyst bed of an adiabatic reactor with a heat source,wherein the catalyst bed comprises an oxidative coupling of methanecatalyst; (b) introducing a gaseous feed stream comprising methane (CH₄)and oxygen (O₂) having a temperature of less than 350° C. and an initialCH₄:O₂ molar ratio to the adiabatic reactor; (c) igniting the oxidativecoupling of methane reaction; and (d) incrementally reducing the initialCH₄:O₂ molar ratio of the gaseous feed stream introduced into thereactor to a final CH₄:O₂ molar ratio of 9:1 to 3:1 over steps (a)through (c). Embodiment 14 is the method of embodiment 13, wherein theinitial CH₄:O₂ molar ratio is 5:1 to 40:1 and the final CH₄:O₂ molarratio is 9:1 to 3:1, 8:1 to 5:1, or 6:1 to 5:1, preferably 5.5:1 afterincremental reduction. Embodiment 15 is the method of embodiment 13,wherein the temperature of the gaseous feed stream introduced into thereactor is 150° C. or less and the catalyst bed is preheated to 400° C.to 700° C., preferably 500° C. to 600° C. Embodiment 16 is the method ofembodiment 13, wherein the catalyst bed is preheated to at least theignition temperature of the oxidative coupling of methane reaction atthe initial CH₄:O₂ molar ratio and the reactor is in an ignitedcondition. Embodiment 17 is the method of embodiment 16, wherein ignitedconditions comprise a Zeldovich number (B) of greater than or equal tofour and the product of a Damköhler number (Da) and Zeldovich number(B×Da) of greater than or equal to one. Embodiment 18 is the method ofany one of embodiments 13 to 17, wherein the heat source is removed inany one of steps (a)-(d) after the catalyst bed has been preheated.Embodiment 19 is the method of any one of embodiments 13 to 18, whereinof the characteristic heat removal (heat loss) time from the adiabaticreactor is higher than the residence time required to achievesubstantially complete conversion of the oxygen. Embodiment 20 is themethod of any one of embodiments 13 to 19, continuing the oxidativecoupling of methane reaction after step (d) to produce a product streamcomprising C₂+ hydrocarbons.

The following includes definitions of various terms and phrases usedthroughout this specification.

The phrase “ambient temperature” when used in the context of a reactantfeed during steady state operation of an OCM reaction means atemperature of 40° C. or less, preferably less than 30° C.

The phrase “steady state” refers to a reactor system where theparameters of the reactor system are a constant or substantiallyconstant value. Steady state parameters of the system include reactantfeed temperature, pressure, feed composition including CH₄:O₂ moleratio, feed flow rate, methane and oxygen conversion, productcomposition, and catalyst bed temperature.

The terms “about” or “approximately” are defined as being close to asunderstood by one of ordinary skill in the art. In one non-limitingembodiment the terms are defined to be within 10%, preferably within 5%,more preferably within 1%, and most preferably within 0.5%.

The term “C_(x)+ hydrocarbons” where x is an integer refers to a mixtureof hydrocarbons having a carbon number of x and more. For example C₂+hydrocarbons is a mixture of hydrocarbons having 2 and more carbonnumbers.

The term “substantially” and its variations are defined as being largelybut not necessarily wholly what is specified as understood by one ofordinary skill in the art. In one non-limiting embodiment, substantiallyrefers to ranges within 10%, within 5%, within 1%, or within 0.5%.

The terms “wt. %”, “vol. %”, mol. % refers to a weight, volume, or molarpercentage of a component, respectively, based on the total weight, thetotal volume, or total moles of material that includes the component. Ina non-limiting example, 10 grams of component in 100 grams of thematerial is 10 wt. % of component.

The terms “inhibiting” or “reducing” or “preventing” or “avoiding” orany variation of these terms, when used in the claims and/or thespecification, include any measurable decrease or complete inhibition toachieve a desired result.

The term “effective,” as that term is used in the specification and/orclaims, means adequate to accomplish a desired, expected, or intendedresult.

The use of the words “a” or “an” when used in conjunction with any ofthe terms “comprising,” “including,” “containing,” or “having” in theclaims or the specification may mean “one,” but it is also consistentwith the meaning of “one or more,” “at least one,” and “one or more thanone.”

The words “comprising” (and any form of comprising, such as “comprise”and “comprises”), “having” (and any form of having, such as “have” and“has”), “including” (and any form of including, such as “includes” and“include”), or “containing” (and any form of containing, such as“contains” and “contain”) are inclusive or open-ended and do not excludeadditional, unrecited elements or method steps.

The methods of the present invention can “comprise,” “consistessentially of,” or “consist of” particular ingredients, components,compositions, etc., disclosed throughout the specification. With respectto the transitional phase “consisting essentially of,” in onenon-limiting aspect, a basic and novel characteristic of the methods ofthe present invention is the ability to provide the gaseous feed streamto the reactor at ambient temperature during steady state operation ofan OCM reaction.

Other objects, features and advantages of the present invention willbecome apparent from the following figures, detailed description, andexamples. It should be understood, however, that the figures, detaileddescription, and examples, while indicating specific embodiments of theinvention, are given by way of illustration only and are not meant to belimiting. Additionally, it is contemplated that changes andmodifications within the spirit and scope of the invention will becomeapparent to those skilled in the art from this detailed description.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1A shows a schematic representation of an oxidative coupling ofmethane steady state reactor exit temperature as a function of feedtemperature, and oxygen to methane feed ratio.

FIG. 1B shows a schematic of the projection of the ignition andextinction locus in the plane of oxygen to methane ratio and feed gastemperature.

FIG. 1C shows another schematic representation of an oxidative couplingof methane steady state reactor exit temperature as a function of feedtemperature, and oxygen to methane feed ratio.

FIGS. 2A and 2B are a schematic of embodiments of a system to produceC₂+ hydrocarbons using the method of the present invention. FIG. 2Aincludes a single gaseous feed stream and FIG. 2B includes two gaseousfeed stream inlets.

FIG. 3 is graphical representation of CH₄ conversion and C₂+ selectivityof a La₂O₃/CeO₂ catalyst driven OCM reaction, with a La/Ce weight ratioof about 15, a feed molar CH₄:O₂ of 4, a residence time of about 7milliseconds, and a feed temperature of near ambient.

FIG. 4 is graphical representation of ignition and extinction of aLa₂O₃/CeO₂ catalyst driven OCM reaction, with a La/Ce weight ratio ofabout 15, in a 4 mm and a 10.5 mm I.D. reactor.

While the invention is susceptible to various modifications andalternative forms, specific embodiments thereof are shown by way ofexample in the drawings and may herein be described in detail. Thedrawings may not be to scale.

DETAILED DESCRIPTION OF THE INVENTION

Embodiments of the invention involve the use of parameters of the feedgas to an oxidative coupling of methane reactor to establish andmaintain the oxidative coupling of methane reaction in the reactor. Forexample, the OCM reaction is ignited in a manner that avoids transientstate at temperatures that would destroy the catalyst. Once the OCMreaction is in an ignited state, the feed gas may be used as a coolantso that the reactor may be operated in an autothermal and ignited state.In the autothermal state, the oxidative coupling of methane reactionuses only the heat produced by the reaction itself. In other words, inthe autothermal state, no external heating is provided to carry out theoxidative coupling reaction at steady state. In the ignited state, thecatalyst is above its ignition temperature. The catalyst reaches itsignition temperature when the catalyst heats up to a point so that noexternal heating is required. At the ignition temperature, the rate ofheat generation exceeds the rate of heat removal (by the flow). Thereaction is ignited successfully when the catalyst is at, or above, itsignition temperature, but below temperatures at which the catalyst isdestroyed and when the OCM reaction is occurring.

As mentioned previously, the oxidative coupling reaction is soexothermic that the temperature in the reactor may far exceed theignition temperature of the catalyst and reach a temperature thatdestroys (deactivates) the catalyst. Such high temperatures often occurin a transient state. Hence, embodiments of the invention may alsoinclude an innovative start-up procedure that enables the attainment ofan autothermal and ignited state while avoiding a transient state inwhich the temperatures may be high enough to destroy the catalyst. Thestart-up procedure may best be understood by considering how the variousparameters interact with each other to create reaction conditions in theoxidative coupling of methane reactor.

FIGS. 1A to 1C show a range of operating parameters in which both anignited and an extinguished state coexist for the OCM reaction and theregion of steady-state multiplicity. FIG. 1A is a three-dimensionalgraph showing the difference between the reactor exit and feedtemperatures (ΔT) on the y-axis, temperature of the inlet feed (T_(in))on the x-axis and the ratio of O₂/CH₄ in the reactor on the z-axis. FIG.1B is a top plan view of FIG. 1A, which shows T_(in) on the x-axis inrelation to the ratio O₂/CH₄ on the z-axis. FIG. 1C is a cross-sectionalview of FIG. 1A at successively higher values of O₂/CH₄ ratio, whichshows ΔT on the y-axis in relation to the T_(in) on the x-axis. The pathlabeled A-B-C shows one example of a preferred start-up path going frominitial (extinguished) state A to a final ignited operating state C.

In embodiments of the invention, the special start-up procedure may befollowed to enable attainment of a stable ignited state, while avoidingthe problem of transient heating of the catalyst to temperaturessufficient to decompose the catalyst or significantly reduce thecatalyst activity. In embodiments of the invention, this start-upprocedure involves simultaneously changing more than one controlvariable to enable the system to move along a special path from aninitial state to a final ignited operating state, namely from state A,to state C, via state B, as shown in FIGS. 1A to 1C. This ultimately canallow for the use of a reactant feed at substantially reducedtemperatures compared to the adiabatic temperature for a given feedratio (for example a feed temperature 600-900° C. lower than theadiabatic temperature). The feed temperature at point C is less thanthat at point B, which is less than at point A.

Referring to FIGS. 1A and 1B, line E represents the extinction locus ofthe catalyst and line I represents the ignition locus of the catalyst.The catalyst is extinguished when it is no longer able to maintainautothermal operation and the oxygen conversion is very low (e.g.,<20%). The catalyst is ignited when it reaches a temperature at whichthe activity is high enough to maintain autothermal operation with highoxygen conversion (e.g., >80%).

Starting from some low value of T_(in) (anywhere to the left of line E)in FIG. 1B, ignition occurs when the feed temperature is increasedenough to cross line I, the ignition locus, while moving from left toright. Once the reactor is in the ignited state, reducing T_(in) causesa gradual reduction in reactor (catalyst) temperature until line E iscrossed (moving from right to left); at which point the reactionextinguishes and there is an abrupt decrease in catalyst and reactortemperature.

Considering FIG. 1A, once in the ignited state, the region to the rightof line I remains in that state. The region “O,” enclosed by theignition and extinction loci (denoted by curves I and E in FIG. 1B), isthe region of viable performance in the reactor for the oxidativecoupling reaction to produce ethane and/or ethylene. Within this region,operation close to the extinction locus is optimal.

In embodiments of the invention, the oxidative coupling reaction processis designed to take place at state C in region O. Further, inembodiments of the invention the reaction is carried out at steady stateat state C, having moved the reaction conditions from startup conditionA outside of region O, proceeding to intermediate state B and thenfinally to state C.

Referring to FIGS. 1A and 1C, line RT represents a plot of reactor exittemperature (T) versus inlet feed temperature T_(in). Line T_(c) showsthe maximum operating temperature of the catalyst and line T_(o) showsthe operating temperature of the reactor. Temperatures at and aboveT_(c) will destroy or deactivate the catalyst. For example, if thetemperature of the catalyst exceeds 700° C., the catalyst mightevaporate or change its oxidation state. For pelletized catalysts, thethermal stress at such a high temperature may cause the pellets tofracture, disintegrate and get blown out of the reactor.

In embodiments of the invention, the oxidative coupling reaction iscarried out in region O, preferably at or near state C to achievefavorable conditions for conversion to ethane and/or ethylene withoutdestroying the catalyst by excessive heat. To get to state C,embodiments of the invention start at point A (highest inlet feedtemperature (T_(o)), and lowest ratio O₂/CH₄—relative to points B andC). The reaction conditions are changed to proceed along line A-B-C toreach point B (approximately equal exit temperature, lower inlet feedtemperature (T_(o)), and higher ratio O₂/CH₄—relative to point A). Thereaction conditions are further changed to proceed along line A-B-C toreach point C (approximately equal exit temperature, lowest inlet feedtemperature (T_(o)), and highest ratio O₂/CH₄—relative to points A andB). Developing the reaction conditions in this way along line A-B-C, asillustrated, the reactor attains an autothermal state while avoiding atransient state in which the catalyst is exposed to high transienttemperatures, which could potentially extinguish the catalyst. Once thereaction conditions are at state C, the feed to the reactor is at a lowenough temperature to provide a high cooling rate while still enablingoperation in the autothermal state.

Thus, the present application provides for an OCM startup process thatcan result in the use of a gaseous feed stream (e.g., CH₄ and O₂) atambient or near ambient temperatures after the initial startup phase hasbeen completed and the reaction has reached steady state operatingconditions. The startup process can be used with any known OCM catalyst.The startup procedure involves igniting the reaction and then changingat least one operating parameter to obtain desired steady stateconditions and using near ambient reactant feed temperatures to keep thereaction in an auto-thermal state. In embodiments of the invention, suchas some embodiments where a portion of the feed gas is recycled from acryogenic separation tower, the temperature may be well below ambienttemperature for example much below 30° C. In such embodiments, the feedgas rate may be reduced as less feed gas would be required to cool thereactor and maintain it in an auto-thermal state.

In embodiments of the invention, the startup procedure involves startingat a high temperature feed (at state A, FIG. 1) and varying anycombination of the following parameters: (1) ratio of O₂/CH₄, (2)temperature of feed, (3) residence time in the reactor, and (4) initialtemperature of the catalyst.

By implementing the startup procedure and implementing a low temperaturefeed as described herein, embodiments of the invention providesadvantages such as: (1) reducing capital costs associated with largeheat exchangers needed to pre-heat the feed gas to high temperatures forignition; (2) use of an economical adiabatic reactor instead of a cooledmulti-tubular reactor or a more complex fluidized bed reactor; (3)operation with higher methane conversion than is possible with a cooledmulti-tubular or fluidized bed reactor; and (4) stable operation at thehighest possible throughput or production rate leading to a smallerreactor.

These and other non-limiting aspects of the present invention arediscussed in further detail in the following sections.

A. Oxidative Coupling of Methane Process

In an OCM reaction, a gaseous feed mixture containing methane (CH₄) andoxygen (O₂) can be contacted with an OCM catalyst under suitableconditions to produce a product stream that includes C₂+ hydrocarbons.The C₂+ hydrocarbons are obtained from oxidative coupling of CH₄. FIG.2A is a schematic of a reactor system for an OCM process having a singlegaseous feed mixture. Referring to FIG. 2, a schematic of system 10 forthe production of ethylene is depicted. System 10 may include anadiabatic reactor 12 and a catalytic bed 14 that includes catalyticmaterial capable of catalyzing an oxidative coupling of methanereaction. Gaseous methane stream 16 and gaseous oxygen source stream 18can be combined to produce gaseous feed stream 20. Flow controller 22can be used to control the amount of methane and/or oxygen fed toreactor 12. As shown, flow controller 22 is a three-way valve, however,any type of flow controller (e.g., mass flow meter) or valveconfiguration or series of flow controllers or valves can be used tocontrol the flow of the oxygen source and methane source to reactor 12.In some embodiments, the methane source and the oxygen source are fed asindividual gas feeds to reactor 12, and form the gaseous feed mixture 20in reactor 12. Referring to FIG. 2B, methane feed stream can enterreactor 12 through inlet 30. Oxygen source feed stream 18 can enterreactor 12 through inlet 32. Inlets 30 and 32 can include necessaryvalves or meters to control the flow and/or amount of gas enter thereactor 12 for both methane feed stream 16 and oxygen source feed stream18. The reactor 12 can include other conventional components forcontrolling chemical reactions such as, for example, heating elements,thermocouples, manual and/or automated controllers, valves, and thelike. Average operating pressure of reactor 12 can be about 0.1 MPa,with other ranges contemplated 0.1 to 0.2 MPa, preferably between 0.2and 1.0 MPa.

The following subsections provide non-limiting startup procedures forinitiating an OCM reaction. In addition to being able to use a reactantfeed at or near ambient temperatures during steady state operation,these startup procedures can advantageously limit excessive heatproduction. Wishing not to be bound by theory, it is believed thatlimiting excessive heat production can result in lower carbon dioxideand/or carbon monoxide formation during the OCM reaction, therebyincreasing C₂+ selectivity.

1. Incremental Change of CH₄:O₂ Molar Ratio and Feed Temperature

The startup of the reactor of the present invention can be controlledsuch that the sum of the gaseous feed stream temperature and a rise of atemperature of the reactor over a startup period is close to a finalcatalyst operating temperature of the reaction. Such control allows thereaction to be performed in adiabatic reactors instead of cooledmulti-tubular reactors or fluidized bed reactors. During startup,gaseous feed stream 20 can have an initial CH₄:O₂ molar ratio of 5:1 to40:1, or 8:1 to 40:1, or 10:1 to 40:1, or 15:1 to 35:1, or 20:1 to 30:1,or any range or value there between (e.g., 10:1, 11:1, 12:1, 13:1, 14:1,15:1, 16:1, 17:1, 18:1, 19:1, 20:1, 21:1, 22:1, 23:1, 24:1, 25:1, 26:1,27:1, 28:1, 29:1, 30:1, 31:1, 32:1, 33:1, 34:1, 35:1, 36:1, 37:1, 38:1,39:1, 40:1). In some embodiments, gaseous feed stream 20 or theindividual streams (streams 16 or 18) that make up the gaseous feedstream can be at a temperature considered to be below or at ambienttemperatures of an OCM process prior to entering reactor 12. By way ofexample, gaseous feed stream 20 can be made up of recycled streams froma cryogenic separation process. In some embodiments, gaseous feed stream20 can have a temperature of less than 400° C.

Referring to FIG. 2A, gaseous feed stream 20 can be heated by passingthe stream through preheater 24 to heat the stream to a temperature ofat least 400° C., or 400° C. to 750° C., or 550° C. to 650° C., or 500°C. to 600° C., or any value or range there between (e.g., 400° C., 425°C., 450° C., 475° C., 500° C., 525° C., 550° C., 575° C., 600° C., 625°C., 650° C., 675° C., 700° C., or 725° C., 750° C.). As the preheatedfeed passes over the catalyst bed, ignition of the catalyst can occurand the oxidative reaction commences (See reaction equations (V) through(VIII)). In some embodiments, gas stream 26, which is more reactive thanmethane, can be added to gaseous feed stream 20 prior to reactionignition. After ignition, flow of gas stream 26 can be slowed ordiscontinued by adjusting a flow controller associated with the reactivegas stream. Non-limiting examples of gases more reactive (with oxygen)than methane include hydrogen, ethane, ethylene, propane, propylene,diazomethane, tributylborane, and triethylborane. At this stage of theOCM startup process, the reactor 12 and/or the catalytic bed 14 is notheated. In other instances, however, the reactor 12 and/or catalytic bed14 can be heated.

Upon ignition of the oxidative coupling of methane reaction, the productstream 28, which includes the C₂+ hydrocarbons, is produced. At thistime, both the temperature and the CH₄:O₂ molar ratio of the gaseousfeed stream introduced into the adiabatic reactor can be reduced to atemperature of 10° C. to 350° C. and a final CH₄:O₂ molar ratio of 9:1to 3:1. The temperature of gaseous feed stream 20 can be reduced in 1 to10° C. increments by reducing the heat provided by preheater 24. Forexample, preheater 24 can be turned off or down to allow the temperatureof the gaseous feed stream to decrease in 1° C., 5° C., 10° C., 20° C.increments, or increments of 1° C., 2° C., 3° C., 4° C., 5° C., 6° C.,7° C., 8° C., 9° C., 10° C., 15° C., 20° C., etc. In some embodiments,cooling can be used to control the reduction of the gaseous feed stream.Preheater 24 can be a furnace, heat exchanger, heater, electricalheater, steam or the like. The final catalyst operating temperature,after the initial startup procedures have been performed, can be 750° C.to 1100° C. or 850° C. to 950° C., or 750° C., 760° C., 770° C., 780°C., 790° C., 800° C., 810° C., 820° C., 830° C., 840° C., 850° C., 860°C., 870° C., 880° C., 890° C., 900° C., 910° C., 920° C., 930° C., 940°C., 950° C., 960° C., 970° C., 980° C., 990° C., 1000° C., 1010° C.,1020° C., 1030° C., 1040° C., 1050° C., 1060° C., 1070° C., 1080° C.,1090° C., 1100° C. Using flow controller 22 of FIG. 2A, or flowcontrollers in inlets 30 and 32 of FIG. 2B, the flow of oxygen andmethane to reactor 12 can be adjusted to change the CH₄:O₂ molar ratioin the gaseous feed stream entering reactor 12 in 0.01 to 1 molarincrements (e.g., 0.01, 0.02, 0.05, 1, 1.5, 2, etc., increments), whilemaintaining a desired flow rate. The final CH₄:O₂ molar ratio can be 9:1to 3:1, 8:1 to 5:1, or 6:1 to 5:1, preferably 5.5:1 after incrementalreduction. The incremental reduction rate of the gaseous feed streamtemperature and the CH₄:O₂ molar ratio can be correlated to the rise ofthe temperature of the adiabatic reactor over the startup period (e.g.,1 minute to 5 hours) and of the final catalyst operating temperature, asshown in the following equation (IX).T _(f) +ΔT _(ad) =T _(out)  (IX)where T_(f) is the gaseous reactant feed temperature entering thereactor 12, ΔT_(ad) is the adiabatic temperature rise of the reactor 12,and Tout is the final catalyst or catalyst bed operating temperature atthe exit of the catalyst bed. ΔT_(ad) is directly proportional to theoxygen concentration in the feed (or equivalently the O₂ to CH₄ molarratio).

2. Incremental Change of Residence Time

In combination with the parameters described above, the rate of thegaseous feed stream can be controlled to provide a first residence time(e.g., 0.1 to 1000 milliseconds, or 50 to 900 milliseconds, 100 to 800milliseconds, 200 to 700 milliseconds, etc.) in the catalyst bed beforeand during ignition. Once ignition has commenced, the flow of thegaseous feed stream can be adjusted to afford a smaller residence time(e.g., 20 to 0.1 milliseconds, 15 to 1 milliseconds, or 10 to 2milliseconds) in the catalyst bed as compared to the first residencetime.

3. Incremental Change of CH₄:O₂ Molar Ratio at “Ignition” Conditions

In some embodiments, the startup procedure can include preheating thecatalyst bed to a temperature near or above the ignition conditions. Bypreheating the catalyst bed, the gaseous feed stream can be introducedat a lower temperature but still sufficient for ignition of the catalystbed into the auto-thermal state while avoiding uncontrolled runaway ofthe reaction. Uncontrolled runway conditions are conditions in which thetemperature within the catalyst bed is so high as to damage thecatalyst.

The proper ignition conditions can be determined using a combination ofa Zeldovich number (B) and a Damköhler number (Da). The Zeldovich andDamköhler numbers are dimensionless groups. Zeldovich number is thenon-dimensional adiabatic temperature rise given by,B=[(−ΔH)·Y_(O2)/C_(p)]×[E/RT²]; where ΔH is the overall heat (orenthalpy change) of reaction, Y_(O2) is the mole fraction of oxygen inthe reactant mixture, C_(p) is the molar heat capacity of the reactantmixture, E is the activation energy of the reaction, R is the gasconstant, and T is the absolute feed temperature. The Damköhler numberis the non-dimensional residence time given by, Da=k(T)·τ; where k(T) isthe overall (first order) rate constant for the reaction evaluated atthe feed temperature and τ is the residence time in the catalyst bedevaluated at the feed temperature.

There are two conditions necessary for operation in the autothermal (orignited) state. First, steady state multiplicity must be possible forsome combinations of the control variables; the condition for this isBA. Second, the reactor must be started up in a manner such that thecondition for ignition is exceeded; the product of Damköhler number (Da)and Zeldovich number (B×Da) is greater than or equal to one. (BDa≥1) isthe condition for ignition in an adiabatic reactor.

Processes according to embodiments of the invention can includeintroducing gaseous feed stream 20 with an initial CH₄:O₂ molar ratio atas low a residence time as possible to catalyst bed 14. The catalyst bed14 can be preheated to an ignition temperature. By way of example,catalyst bed 14 can be preheated to 400° C. to 700° C., preferably 550°C. to 650° C., or 400° C., 410° C., 420° C., 430° C., 440° C., 450° C.,460° C., 470° C., 480° C., 490° C., 500° C., 510° C., 520° C., 530° C.,540° C., 550° C., 560° C., 570° C., 580° C., 590° C., 600° C., 610° C.,620° C., 630° C., 640° C., 650° C., 660° C., 670° C., 680° C., 690° C.,700° C. The residence time of gaseous feed stream 20 in catalyst bed 14can be 0.1 to 20 milliseconds, 1 to 15 milliseconds, or 2 to 10milliseconds, or 0.1, 0.5, 1, 1.5, 2, 2.5, 3, 3.5, 4, 4.5, 5, 5.5, 6,6.5, 7, 7.5, 8, 8.5, 9, 9.5, 10, 10.5, 11, 11.5, 12, 12.5, 13, 13.5, 14,14.5, 15, 15.5, 16, 16.5, 17, 17.5, 18, 18.5, 19, 19.5, or 20. Gaseousfeed stream 20 can have a temperature of less than 350° C., less than150° C., or 15 to 350° C., or 100 to 300° C., 125 to 250° C. The feedstream 20 does not have to be subjected to any additional heating priorto entering reactor 12 (e.g., preheater 24 does not have to be used). Asgaseous feed stream 20 enters heated catalyst bed 14, the OCM reactioncan ignite. After ignition of the OCM reaction, the initial CH₄:O₂ molarratio of gaseous feed stream 20 can be incrementally reduced to a finalCH₄:O₂ molar ratio of 9:1 to 3:1 (e.g., 8:1 to 5:1, or 6:1 to 5:1,preferably 5.5:1) by adjusting the methane content and oxygen content inthe gaseous feed stream (e.g., adjusting the flow of gaseous methanestream 16 and/or oxygen feed stream 18 using valve 22). In someembodiments, after catalyst bed 14 has been preheated, the heating canbe turned down or off during introduction of gaseous feed stream 20and/or ignition of the OCM reaction. In certain embodiments, the heatgenerated is sufficient to maintain or increase the temperature of thecatalyst bed.

4. Continued Production of C₂+ Hydrocarbons

In the methods described above, after startup is complete (e.g., whenthe reaction reaches steady state conditions) the OCM reaction iscontinued by maintaining the reaction in an autothermal state usinggaseous methane stream 16 as a coolant while product stream 26 iscontinued to be produced. Steady state conditions can include a constantreaction temperature (e.g., the reaction temperature does not vary bymore than 10% from a selected reaction temperature—by way of example, aselected reaction temperature of 850° C. can include a temperature range765° C. to 935° C.). Product stream 26 can exit adiabatic reactor 12 andbe collected, stored, transported, or processed into other chemicalproducts. By way of example, product stream 26 that includes C₂+hydrocarbons and water produced from the reaction can be collected in acollection device and/or transported via piping to a separation unit. Inthe separation unit, the C₂+ hydrocarbons can be separated using knownseparation techniques, for example, distillation, absorption, membranetechnology, etc., to produce an ethylene containing product. Inembodiments when carbon dioxide is in the reactant mixture and/orgenerated in situ, the resulting gases (for example, CO, H₂, andethylene) produced from the systems of the invention (for example,system 10) is separated from the hydrogen, carbon monoxide, and carbondioxide (if present) using known separation techniques, for example, ahydrogen selective membrane, a carbon monoxide selective membrane, acarbon dioxide selective membrane, or cryogenic distillation to produceone or more products such as ethylene, carbon monoxide, carbon dioxide,hydrogen, or mixtures thereof. The products can be used in additionaldownstream reaction schemes to create additional products or for energyproduction. Examples of other products include chemical products such asmethanol production, olefin synthesis (e.g., via Fischer-Tropschreaction), aromatics production, carbonylation of methanol,carbonylation of olefins, the reduction of iron oxide in steelproduction, etc. The method can further include isolating and/or storingthe produced gaseous mixture or the separated products.

B. Catalytic Material and Reactants

One or more OCM catalysts can be used in the process of the presentinvention. The catalyst(s) can be a supported catalyst(s), a bulk metalcatalyst(s), or an unsupported catalyst(s). The support can be active orinactive. The catalyst support can include MgO, Al₂O₃, SiO₂, or thelike. All of the support materials can be purchased or be made byprocesses known to those of ordinary skill in the art (e.g.,precipitation/co-precipitation, sol-gel, templates/surface derivatizedmetal oxides synthesis, solid-state synthesis, of mixed metal oxides,microemulsion technique, solvothermal, sonochemical, combustionsynthesis, etc.). One or more of the catalysts can include one or moremetals or metal compounds thereof. Non-limiting catalytic metals includeLi, Na, Ca, Cs, Mg, La, Ce, W, Mn, Ru, Rh, Ni, or Pt, or combinations oralloys thereof. Non-limiting examples of catalysts of the inventioninclude: (1) La on a MgO support; (2) Na, Mn, and La₂O₃ on an aluminumsupport; (3) Na and Mn oxides on a silicon dioxide support; (4) Na₂WO₄and Mn on a silicon dioxide support, or any combination thereof.Non-limiting examples of some particular catalysts that can be used inthe context of the present invention to promote oxidative coupling ofmethane to produce ethylene are Li₂O, Na₂O, Cs₂O, MgO, WO₃, Mn₃O₄,La₂O₃/CeO₂, SrO/La₂O₃, Yb₂O₃—SrO—CeO₂, Li/MgO, Na₂WO₄—Mn—O/SiO₂, or anycombination thereof.

The gaseous feeds stream in the context of the present invention can bea gaseous mixture that includes, but is not limited to, a hydrocarbon ormixtures of hydrocarbons and oxygen. The hydrocarbon or mixtures ofhydrocarbons can include natural gas, liquefied petroleum gas containingof C₂-C₅ hydrocarbons, C₆+ heavy hydrocarbons (e.g., C₆ to C₂₄hydrocarbons such as diesel fuel, jet fuel, gasoline, tars, kerosene,etc.), oxygenated hydrocarbons, and/or biodiesel, alcohols, or dimethylether. In a preferred aspect, the hydrocarbon is a mixture ofhydrocarbons that is predominately methane (e.g., natural gas). Theoxygen containing gas used in the present invention can be air, oxygenenriched air, oxygen gas, and can be obtained from various sources. Thereactant mixture may further contain other gases, provided that these donot negatively affect the reaction. Examples of such other gases includecarbon dioxide, nitrogen, and/or hydrogen. The hydrogen may be fromvarious sources, including streams coming from other chemical processes,like ethane cracking, methanol synthesis, or conversion of methane toaromatics. Carbon dioxide may be from natural gas, or a waste or recyclegas stream (e.g., from a plant on the same site, like for example fromammonia synthesis) or after recovering the carbon dioxide from a gasstream.

EXAMPLES

The present invention will be described in greater detail by way ofspecific examples. The following examples are offered for illustrativepurposes only, and are not intended to limit the invention in anymanner.

Example 1 Incremental Change of CH₄:O₂ Molar Ratio and Feed Temperaturewith a La₂O₃/CeO₂ Catalyst

A 10.5 mm I.D. quartz reactor was used as the adiabatic reactor. Agaseous feed mixture that included reactant gases CH₄ and O₂ at a CH₄:O₂molar ratio of 20:1 was introduced to the adiabatic reactor. The gaseousmixture was preheated to about 550-600° C. and had a residence time offrom about 0.1 milliseconds to about 100 milliseconds in the catalystbed that included a La₂O₃/CeO₂ catalyst having a La/Ce wt. ratio of 15.The feed composition and feed temperature during start up were changedin small steps simultaneously to a final CH₄:O₂ molar ratio of 4 to 5,and to an ambient feed temperature (e.g., less than 20° C.). The outletgas from the reactor was determined by GC analysis to include C₂ andhigher hydrocarbons and syngas composition, such as C₂H₄, C₂H₆, CH₄, CO,H₂, CO₂ and H₂O. A steady performance for the duration of experiment(about 20 hours) was achieved. The selectivity to C₂ and higherhydrocarbons was comparable or better than that of conditions of thecomparative examples discussed in Example 2. FIG. 3 shows theperformance (CH₄ and C₂+ selectivity of La₂O₃/CeO₂ catalyst at a finalCH₄:O₂ molar ratio, residence time of about 7 milliseconds; and a feedtemperature equal to near ambient temperature.

Example 2 Comparative Example of Example 1—No Change in StartupProcedures

As a comparative example, an experiment performed with the same catalystand reactor with the final conditions of Example 1 being used as thestart-up and final conditions. Negligible methane conversion (<1%) wasobserved.

Example 3 Comparative Example of Example 1—Change in Temperature

As a comparative example, an experiment performed with the same catalystand reactor and at same final conditions of Example 1 was performed,except that only one parameter i.e., temperature was varied by keepingfeed ratio constant at 4 and residence time of 7 milliseconds. Thereaction was sustained when the feed temperature was reduced to about160° C., but quickly died upon reducing the temperature to ambient.‘Residence time’ refers to the contact time of the flowing gases (atreaction conditions) in the catalyst bed and is defined as the ratio ofvoid volume in the catalyst bed to the actual volumetric flow rate underreactive conditions.

Comparison of Example 1 to Comparative Examples 2 and 3 demonstratedthat varying two parameters during startup (i.e., change in molar ratioand feed temperature) provided a desired steady state condition andconversion of methane instead of little to no conversion of methane(Example 2) when the reaction was started at the desired operatingconditions and/or extinction of the catalyst as demonstrated in Example3 when only the temperature was changed during startup.

Example 4 Incremental Change of CH₄:O₂ Molar Ratio, Feed Temperature andResidence Time with a Na₂WO₄—Mn—O/SiO₂ Catalyst

A 22 mm I.D. quartz tube was used as the reactor. A gaseous feed mixturethat included reactant gases CH₄ and O₂ at a CH₄:O₂ molar ratio of about16 was introduced to the reactor. The gaseous mixture was preheated toabout 650-700° C. and had a residence time of from about 25 millisecondsto about 200 milliseconds in the catalyst bed that included aNa₂WO₄—Mn—O/SiO₂ catalyst. The feed composition and feed temperatureduring start up were changed in small steps simultaneously to a finalCH₄:O₂ molar ratio of 4 to 5, and a temperature of about 100° C. Theresidence time was lowered to about 40 milliseconds when the temperaturewas about 400° C. and the CH₄:O₂ molar ratio was about 5. The outlet gasfrom the reactor was determined by GC analysis to include C₂ and higherhydrocarbons and syngas composition, such as C₂H₄, C₂H₆, CH₄, CO, H₂,CO₂, and H₂O. A steady performance for the duration of the experimentwas achieved at a final reactor feed temperature of 110° C., a finalCH₄:O₂ molar ratio of about 3.5, and a residence time of 40milliseconds. The performance of the Na₂WO₄—Mn—O/SiO₂ catalyst underthese reaction conditions are presented in Table 1. The percent methaneconversion was 27%.

TABLE 1 C₂+ Selectivity % Ethylene 25.43 % Ethane 9.33 % Propene 2.62 %Propane 0.3 % total C₂+ 37.68 % CO 32.92 % CO₂ 29.39 % H₂/CO 0.82Ethylene/ethane 2.73

Example 5 Comparative Example of Example 4—No Change in StartupProcedures

As a comparative example, an experiment performed with the same catalystand reactor and at same final conditions of Example 4 was performed, butwithout preheating the catalyst or feed gas to sufficiently hightemperature to achieve the ignited state. Negligible methane conversion(<1%) was observed, because only the extinguished state was reached.

Comparison of Example 4 to Comparative Example 5 demonstrated thatvarying two parameters during startup (i.e., change in molar ratio andfeed temperature) provided a desired steady state condition andconversion of methane instead of little to no conversion of methane(Example 2) when the reaction was started at the desired operatingcondition.

Example 6 Incremental Change of CH₄:O₂ Molar Ratio, Feed Temperaturewith a Yb₂O₃—SrO—CeO₂ Catalyst

A 22 mm I.D. quartz tube was used as the reactor. A gaseous feed mixturethat included reactant gases CH₄ and O₂ at a CH₄:O₂ molar ratio of about20 was introduced to the reactor. The gaseous mixture was preheated toabout 600° C. and had a residence time of from about 0.5 milliseconds toabout 5 milliseconds in the catalyst bed that included a Yb₂O₃SrO—CeO₂catalyst. The feed composition and feed temperature during start up werechanged in small steps simultaneously to a final CH₄:O₂ molar ratio ofabout 5, and a temperature of about 130° C. The outlet gas from thereactor was determined by GC analysis to include C₂ and higherhydrocarbons and syngas composition, such as C₂H₄, C₂H₆, CH₄, CO, H₂,CO₂, and H₂O. A steady performance for the duration of the experimentwas achieved at a final feed temperature of 130° C., a final CH₄:O₂molar ratio of about 4.7, and a residence time of 4 milliseconds. Theperformance of the Yb₂O₃—SrO—CeO₂ catalyst under these reactionconditions are presented in Table 2. The percent methane conversion was11.53%.

TABLE 2 C₂+ Selectivity % Ethylene 21.19 % Ethane 14.22 % Propene 0.77 %Propane 0.41 % total C₂+ 36.58 % CO 13.18 % CO₂ 50.24 % H₂/CO 2.32Ethylene/ethane 1.49

Example 7 Comparative Example of Example 6—No Change in StartupProcedures

As a comparative example, an experiment performed with the same catalystand reactor and at final conditions of Example 6 without following thepreferred path. Only an extinguished final state was reached withnegligible methane conversion (<1%).

Comparison of Example 6 to Comparative Example 7 demonstrated thatvarying two parameters during startup (i.e., change in molar ratio andfeed temperature) provided a desired steady state condition andconversion of methane instead of little to no conversion of methane(Example 2) when the reaction was started at the desired operatingcondition.

Example 8 Effect of Reactor Size

The effect of reactor size on heat removal was investigated using 4 mmand 10.5 I.D. reactor tubes by changing only the furnace temperature(i.e., all other parameters were kept constant). The reaction parameterswere a CH₄:O₂ molar ratio of about 4 and a residence time of about 2milliseconds. A La—Ce catalyst having a La/Ce weight ratio of 15 wasused. FIG. 4 shows ignition and extinction of the La—Ce oxide catalystfor the 4 mm and 10.5 mm I.D. reactors, respectively. It was observedwhen a larger reactor size was used, the characteristic heat removaltime was much larger and the extinction temperature was lowered further.From the data obtained from the 10.5 mm reaction (FIG. 4) it wasdetermined that the extinction temperature was lower than 160° C. ascompared to about 270° C. for the smaller reactor. When the furnace wasopened in the larger reactor to further reduce the temperature close toambient, the reactor extinguished. Also, the ignition temperature wasabout 400° C., which was lower than the ignition temperature of 475° C.observed in the smaller reactor.

The invention claimed is:
 1. A startup method for an oxidative couplingof methane reaction to produce C₂+ hydrocarbons, the method comprisingthe steps of: (a) preheating a catalyst bed of an adiabatic reactor witha heat source, wherein the catalyst bed comprises an oxidative couplingof methane catalyst; (b) introducing a gaseous feed stream comprisingmethane (CH₄) and oxygen (O₂) having a temperature of less than 350° C.and an initial CH₄:O₂ molar ratio to the adiabatic reactor; (c) ignitingthe oxidative coupling of methane reaction; and (d) incrementallyreducing the initial CH₄:O₂ molar ratio of the gaseous feed streamintroduced into the reactor to a final CH₄:O₂ molar ratio of 9:1 to 3:1over steps (a) through (c).
 2. The method of claim 1, wherein theinitial CH₄:O₂ molar ratio is 5:1 to 40:1 and the final CH₄:O₂ molarratio is 9:1 to 3:1 after incremental reduction.
 3. The method of claim1, wherein the temperature of the gaseous feed stream introduced intothe reactor is 150° C. or less and the catalyst bed is preheated to 400°C. to 700° C.
 4. The method of claim 1, wherein the catalyst bed ispreheated to at least the ignition temperature of the oxidative couplingof methane reaction at the initial CH₄:O₂ molar ratio and wherein thereactor is in an ignited condition.
 5. The method of claim 4, whereinignited conditions comprise a Zeldovich number (B) of greater than orequal to four and the product of a Damköhler number (Da) and Zeldovichnumber (B×Da) of greater than or equal to one.
 6. The method of claim 1,wherein the heat source is removed in any one of steps (a)-(d) after thecatalyst bed has been preheated.
 7. The method of claim 1, wherein ofthe characteristic heat removal time from the adiabatic reactor ishigher than the residence time required to achieve substantiallycomplete conversion of the oxygen.
 8. The method of claim 1, furthercomprising continuing the oxidative coupling of methane reaction afterstep (d) to produce a product stream comprising C₂+ hydrocarbons.